Hydrorefining of heavy mineral oils



March 27, 1962 c. P. REEG ETAL HYDROREFINING OF HEAVY MINERAL OILS Filed Jan. 27, 1958 Unite States This invention relates to methods for pretreating heavy oils, and especially crude oils, whereby they may be subsequently treated in a catalytic hydrorening operation with `a minimum deposition of deactivating deposits upon the catalyst. ln broad aspect, the invention consists in rst subjecting the feed oil to a combined partial vaporization-deasphalting step in the presence of the hydrogen-containing recycle gas from the hydrorener, and at substantially the pressure prevailing in the hydrorefiner. The asphaltic residue from the partial vaporization step is then subjected to a thermal coking step to produce further quantities of asphalt-free oil in the form` of a coker distillate, the asphaltic bodies being decomposed largely to coke. At least the light portion of the coker distillate is then combined with the gaseous mixture derived from the vaporization step, and the combined product is subjected to hydrorening. The heavy ends of the coker distillate may be recycled to the partial vaporization step to further reduce deposit-forming bodies in the oil.

The principal objective of the invention is to provide means for hydroretining crude oils, or heavy fractions thereof, while avoiding to a maximum extent the normally occurring deposition of coke and other deactivating deposits upon the catalyst. Another objective is to provide methods for removing the precursors of such deactivating deposits from the feed oil, While maintaining maximum over-all liquid yields. Still another object is to remove such deposit-forming components in such manner as to achieve control over the pour point of the final hydrogenated product. Still another object is to achieve maximum vaporization of the feed oil by pressure stripping with the hydrorener recycle gas in such manner as to minimize utility costs. Specically, it is desired to avoid the expensive depressuring and repressuring of recycle gas. Other objects will lbe apparent from the more detailed description which follows.

The process of this invention is adapted mainly for the treatment of total crude oils, or reduced crude oils. lt is especially adapted for the treatment of crude shale oils, i.e. the full-range oil produced directly from the retorting of shale rock. Crude shale oils present difcult problems in rening in that they contain a high proportion of asphaltenes and carboids and up to about 4% of nitrogen in the form of organic nitrogen compounds. In addition, metals may be present in the form of porphyrin-metal complexes or other organo-metallic compounds. Crude shale oils also possess an unusually high pour point. For example, the crude oil from Colorado shale is normally a solid at room temperature, and its pour point may be in excess of 100 F.

From the standpoint of reiining flexibility, the most desirable initial treatment for such oils consists of catalytic hydroretining to decompose nitrogenous compounds and sulfur compounds, to hydrogenate asphaltenes and carboids, to decompose organo-metallic compounds, and in general to improve the color and handleability of the oil for subsequent refining procedures such as cracking, reforming, and the like. The rst major diiculty encountered in attempting to hydrorene such oils consists in the tendency of the oil to deposit deactivating solids upon hot catalytic surfaces when the feedstock is brought into contact therewith. The exact nature of the depositatent forming constituents is not known, but they may comprise any of the above-noted organo-metallic components, asphaltenes, carboids, and/or nitrogenous compounds.

It has been found that conventional pretreatment steps such as vacuum distillation, thermal coking, or liquid phase deasphalting using lower parains, do not entirely solve the problem of removing such deposit-forming materials, unless severe pretreating conditions are employed entailing sacrificial yields. While the light fractions frorn thermal coking do not cause difficulty, the heavier Coker distillates still tend to form deposits, as does the ratiinate from conventional deasphalting using for example n-pentane as the deasphalting agent. As a corollary to the above, thermal coking and liquid phase deasphalting are subject to the disadvantage that relatively low liquid yields are obtained if those treatments are made sufficiently severe or exhaustive to give a nondeposit-forming distillate or rainate. It is hence highly desirable to provide a more selective method for removing the deposit-forming constituents without removing all of the heavier components, some of which do not tend to form deposits.

In the case of vacuum distillation, it has been found that only about Lt0-50% by volume of non-deposit-forming distillate can be economically produced without employing cracking temperatures. The residue, constituting the vacuum distillation bottoms, must then be subjected to coking or `deasphalting with attendant disadvantages,

above noted. By going to extremely high vacuum, higher distillate yields may be obtained, but such distillates then will include some of the heavier, deposit-forming bodies. The process of this invention is designed to overcome the major disadvantages of each of the foregoing processes.

While it has been indicated that crude oils, and especially shale oils, are the feeds of major concern herein, it is not intended to preclude the use of heavier distillate fractions, some of which present similar problems in catalytic hydrorefining, though to a decidedly lesser extent. In general, any crude oil, heavy distillate, or other heavy fraction having an end boiling point in excess of about 700 F. may be employed. It is also contemplated that desirable, for reasons of economy, to transport the shale' oil after a minimum of processing at the retorting site, to a more distant refinery for final treatment. The most economical form of transport is by pipeline, but this is practical only if the oil has a suiciently low pour point to ow freely during cold winter months. It is therefore desirable to reduce the pour point of the crude oil during the initial refining steps so that it may be economically transported to a more favorably situated refining site. This usually involves reducing the pour point from its initial Sil- F. to about 20-70 F. Such pour point control is integrated into the combined treating process herein described.

The crude shale oil described herein may be produced for example by the retorting procedures described in U.S. Patents 2,501,153, 2,640,015, and 2,640,019. in general, the retorting procedure consists in initiating combustion in a body of moving shale rock, and utilizing the hot combustion gases to educt the oil from uncombusted shale rock upstreamwardly.

The process of this invention will now be described in more detail in connection with the attached ow sheet, which is intended merely to illustrate the principal modifications, but is not intended to be limiting in scope. The

initial feedstock is brought in through line l, and, according to one modification, all or a. portion thereof is diverted through line 2 wherein it may be mingled with a coker distillate bottoms fraction from line' to bedescribed hereinafter, and with the major portion of hydrorefiner recycle gas from line 4. This combined stream is then preheated to a temperature suflicient to achieve the desired vaporization of oil in preheater 6. This temperature may range anywhere between about 200 and 850 F. but is preferably within about 100F. of the temperature to be employed in the subsequent hydrofining step.

Alternatively, the initial feed from line l may be diverted in whole or in part to distillation column 28 via line 27. In distillation column 28 a light asphalt-free fraction of feed, preferably boiling below about 600 F., may be taken overhead via line 29 and mingled directly with the feed to the hydrorener in line d5. This reduces the load on separator 8. Also, the resulting bottoms fraction, which is removed via line 3G, may be diverted in part to coker 24 via line Z2. The proportion so diverted depends upon the relative treating capacities of coker 2Liand separator 8. Ordinarily, however, it is preferable to send all or a major portion of the bottoms from line 30 directly to separator 8 via lilies 31 and 2, and preheater 6.

The mixed phase effluent from preheater 6 is then transferred via line 7 to a centrifugal separator 8, which may be of the Webre cyclone type. This unit consists ofv an outer cylindrical shell 9 enclosing an axially positioned open-ended vapor outlet conduit 10 which extends upwardly to near the top of shell 9', and terminates at its lower end in vapor outlet port l2, located below the ceuter portion of shell 9. The feed inlet line l enters shell l tangentially thereto at a point approximately midway between the top of vapor outlet conduit lil' and the vapor outlet port l2. The liquid portion of entering feed spirals downwardly inside shell 9 by centrifugal force, and forms a liquid level indicated approximately at i3. The entering gas phase spirals upwardly around vapor outlet conduit lil and enters the top thereof and thence 'Hows downwardly and out through vapor outlet port l2. in order to prevent entrainment of liquid in the exiting gas phase, a depending cylindrical drip ring i7 is provided in the top of separator 8. ln this manner, the liquid and gas phases are rapidly and effectively separated, and maximum vaporization, consistent with the desired pourpoint control, is obtained. An additional stream of re cycle L'as may if desired be diverted through line l5, preheated in heater 16 to a temperature 50 to 100 F. in excess or" that prevailing in separator S, and then admitted via line i7 into the lower extremity of separator 8. By this means, additional vaporization by stripping is obtained. Any fresh makeup hydrogen required in the hydrorening operation is preferably admitted via line 18 at this point in order to utilize to a maximum extent its stripping capacity.

The pressure in separator S is preferably maintained' at substantially that prevailing in the hydrorefiner 40. Specifically, in order to obtain maximum deasphalting action and minimum utility costs, it is preferred that the pressure should be within about 300 p.s.i.g. above or below the hydrorener pressure, but any other operative pressure may be used.

The liquid phase accumulating in separator 8 is withdrawn via valve 2li; and line 2l more or less continuously in response to liquid level controller 719. The proportion of feed which is vaporized in separator 8 may be controlled in response to any desired process variable such as temperature, or relative recycle gas Volume admitted thereto. In the modification illustrated, the feed temperature is the controlling variable; when the flow of liquid in line 21 becomes too great, indicating insufficient vaporization, flow-controller 23 opens valve 26 to 4 admit additional heating agent to feed pheheater 6. Too small a flow rate in line 2l, indicating excessive vaporization, is corrected by a reduction in temperature.

Ordinarily, the liquid stream removed in line 21 will comprise only about l-25% by volume of the total liquid feed. It is an important feature of this invention that the 'i5-99% vapor phase fraction removed from separator 8 is produced under deasphalting conditions substantially identical to those prevailin0 in the hydrorener, hence totally excluding therefrom any deposit-forming constituents. To the best of our knowledge this objective cannot be achieved by other methods such as vacuum distillation, coking, or conventional liquid phase deasphalting. The overhead or raffinate of analogous proportions from any such conventional procedures still may contain deposit-forming constituents' which would rapidly deactivate the catalyst in the hydroreliner. The improved results of this invention are achieved moreover with a minimum of utility requirements, in that the major heat input is utilized in the subsequent hydrorefining step, and there is little pressure drop in the system.

The liquid phase withdrawn through line 21 will normally consist of viscous heavy ends together with suspend-ed asphaltenes and carboids precipitated by the deasphalting action of the gas phase iu separator S. rThis liquid phase is then transferred via line 2l to delayed thermal coker 24, together with any fresh feed bottoms from line 22. If necessary, the coker feed may be preheated in heater Z5 to the desired coking temperature.

The character of the coking carried out in delayed coker 24 is conventional in nature. Such operations are generally carried out to reduce the carbon content of low quality crude oils, distillation residues, and sometimes cracking residues. The coking step generally involves heating the oil to a cracking temperature of e.g. 750-950 F. and then transferring the heated oil to the coking drum Where it is allowed to soak in its 0WD heat for several minutes or hours While continuously removing overhead the volatile products boiling in the gas oil range and below, and continuously precipitating coke on the walls of the coking vessel. Ordinarily, none of the liquid feed is removed as liquid, all of the feed being converted either to coke, light gases, light distillates or gas oil.

Eventually the coke which precipitates in coker 24 will fill the vessel, and somewhat before this pointv is reached the operation must be suspended while the coke is removed. As the coke precipitates it adheres to the walls and becomes strongly agglomerated into a dense hard mass. This mass is ordinarily removed either by drilling out the drum, or by washing it out with high-pressure jets of water. While this operation is proceeding, the flow of oil in line 22 is diverted to another coking drum, not shown, where coking proceeds as above described.

Other coking procedures may he employed instead of the delayed thermal coking described above. For example, fluidized coking, or contact coking with a moving bed of coke pebbles, may be employed. These operations are conventional and hence need not be described in detail.

The overhead from coking drum 24 is withdrawn continuously via line 32 and transferred via line 33 to a stabilizer column 34 from which light ends are removed overhead via line 35. The subsequent disposal of the heavier ends from column 34 may vary depending upon feed characteristics, product specifications, and process requirements.

In all cases, it is found that a substantial portion of the intermediate-boiling-range products from column 34 may be subjected directly to hydrorening without deleteriously affecting the catalyst. This is illustrated by the withdrawal of a side-cut through line 37 which may comprise all of the heavier materials boiling below about 700 F. In cases where the heavy ends boiling above 700 F. `are deleterious in the hydrorener, it is prefer-V able to use only the side-cut for hydroreining, and to recycle the entire 700 F.+ bottoms in line 38 via lines 3 and 2 for further deasphalting treatment in separator 8 and coker 24.

When the coker distillate does not contain catalyst deactivating constituents, or contains only a negligible proportion thereof, it will be preferable to eliminate side-cut line 37, and withdraw the total bottoms product through line 38. This may then be diverted in its entirety through lines 39 and 45 to hydroreiiner 40, along with the vapor phase from separator 8, which is being transferred to the hydroreliner via line 45 and preheater 46.

In hydrorener 40 the pretreated mixed feed from separator 8 and column 34 is contacted with a suitable sulfactive hydrorelining catalyst under conditions of hydrorening. The catalyst may be disposed in a fixed stationary bed, or the various moving bed, or fluidized bed techniques may be employed. Generally, the fixed bed technique is most satisfactory. The catalyst may comprise any of the oxides and/ or suliides of the transitional metals, and especially an oxide or sulfide of a group VIII metal (particularly iron, cobalt or nickel) mixed with an oxide or sulfide of a group VIB metal (preferably molybdenum or tungsten). Such catalysts may be employed in undiluted form, but preferably are distended and supported on an adsorbent carrier in proportions ranging between about 2% and 25% by weight. Suitable carriers include in general the dicultly reducible inorganic oxides, eg. alumina, silica, zirconia, titania, clays such as bauxite, bentonite, etc. Preferably the carrier should display little or no cracking activity, and hence highly acidic carriers are generally to be avoided. The preferred carrier is activated alumina, and especially activated alumina containing about 3-15% by weight of coprecipitated silica gel.

The preferred hydrorening catalyst consists of cobalt oxide plus molybdenum oxide supported on silica-stabilized alumina. Compositions containing between about 2% and 8% of CoO, 4% and 20% of M003, 3% and of SiOZ, and the balance A1203, and wherein the mole-ratio of COO/M003 is between about 0.2 and 4, are specifically preferred. These catalysts are preferably prepared by alternate impregnation with aqueous solutions of ammonium molybdate and cobalt nitrate, as described in U.S. Patent No. 2,687,381.

Suitable hydroreiining conditions are as follows:

Within the above operating conditions, the specilic hydrorefining conditions selected should be such as to meet required product specications. Where pour point of the linal product is a critical specification, it is preferable not to employ conditions in hydrorefiner 40 which are severe enough to effect a substantial reduction in pour point. To do so entails the use of severe conditions of temperature and/ or space velocity, such as would tend to decrease liquid yields and increase the rate of deactivation of the catalyst. To meet pour point specifications, it is preferred to vary the conditions in separator S so as to divert a greater or lesser amount of liquid bottoms therefrom to coker 24. Where very viscous feeds are used, and a reduction in pour point is desired, it may be preferable to divert up to about 50% of the feed to coker 24 as bottoms from separator 8. On the other hand, where pour point is not important, or where the feed is initially not excessively viscous, the amount of liquid bottoms withdrawn from separator 8 depends entirely upon the requirement of removing sufficient of the above-notecl components which cause catalyst deactivation. `For this purpose alone it was found that only about 145% of the feed oil ordinarily need be subjected to coking. It will be understood that the evaporation which occurs in separator 8 takes place under noncracking conditions, and hence where the Very heavy ends are evaporated therein, the product may not meet pour point specifications. The major viscosity-reducing operation consists of the coking in coker 24. The desired viscosity of product from hydrorefiner 40 is hence obtained principally by varying the ratio of feed which is vaporized in separator 8 and in coker 24. In other words, the ratio of liquid phase/vapor phase produced in separator 8 is inversely proportional to the desired pour-point of the product.

The product from hydroreiiner 40 is withdrawn via line 42 and condensed in cooler '43 and transferred to high-pressure separator `44. Hydrogen-rich recycle gas, containing minor amounts of light hydrocarbon gases, is withdrawn via line `4,8, and suicient thereof is diverted through lines 4 and 15 to separator 8 to achieve the ends above described. Any remaining portion of recycle gas not utilized in separator 8 may be recycled directly via lines 49 and 45.

The liquid product accumulating in separa-tor `44 may then be `flashed into lowpressure separator 50 via line S1. Low-pressure olf-gas is withdrawn via line 52, and liquid product via line 53. This liquid product is then suitable for use in conventional rening units as for example distillation columns, catalytic cracking units, cata- -lytic reforming units, catalytic hydrocracking, and the like.

The :following example is cited to illustrate one practical application of the process to a crude Colorado shale oil, but `manifestly it is not intended that this should be limiting in scope.

Example Feed charactermica-Gravity, API-204; total nitrogen, 1.84 wt. percent; sulfur, 0.72 wt. percent; Ramsbottom carbon residue, 3 wt. percent; pour point, 90 F.;

initial boiling point-346 F.; 60% lboiling point- Hydrorejmng catalyst- 3% COO plus 9% M003 supported on 1/4 pellets of silica-stabilized alumina (5% SM2-95% A1203), prepared by impregnating pelleted carrier tirst with aqueous ammonium molybdate, then with aqueous cobalt nitrate, drying and calcining.

The total feed plus recycle coker distillate bottoms derived as described below is blended with 3,000 s.c.f./ bbl. of hydrogen-rich recycle gas, and introduced into cyclone separator 8 at a temperature of 780 F. and pressure 2,200 p.s.i.g. An additional stream of fresh makeup hydrogen is introduced via line 17 at 820 F., and at the rate of 1,000 scf/bbl. of initial feed. Liquid bottoms product is withdrawn at the rate of 20 barrels per barrels of combined feed.

T-he bottoms product is then subjected to delayed thermal coking at 35 p.s.i.g. and 815 F. to produce 27% by weight of coke. Based on fresh feed, the coke yield is about 5.4% by weight, and about 70% by volume of the feed to the Coker is recovered as (24+ distillate. The distillate is then fractionated to recover as bottoms an 8% fraction boiling in excess of about 675 F., which is recycled to cyclone separator 8 in admixture with the fresh feed.

The vapor phase from separator 5 is then blended with the overhead from the coker distillate fractionation step. The blend is then preheated to about 700 F. and subjected to hydrorening at an average bed temperature of 800 F., pressure 2,000 p.s.i.g., hydrogen ratio 5,000 scf/bbl., and liquid hourly space velocity 1.0. The product is recovered in 106% yield, contains about 0.11% nitrogen and 0.037% sulfur, has a gravity of 37.3 API, and a pour point of abo-ut 65 F.

Operation under the above conditions can be conspa/,317

tinued for several weeks without shutdown for catalyst regeneration, and without substantial decline in product quality. When the raw crude oil is used directly in the hydrorener, plugging of the reactor and catalyst deactivation require shutdown within a `few hours, or days at the most. When the feed to the hydrorener consists of a 94% fraction of the crude oil, obtained by vacuum distillation with coking of the bottoms, the catalyst life is improved as compared to using the crude oil, but not to the extent obtained when using the 94% feed fraction prepared as described in the foregoing example.

The foregoing description of specilic methods and materials yfor use in this invention is not intended to be limiting in scope except where indicated. Many variations wi-ll occur to those skilled in the art and all such variations which yield essentially the same result are intended to be included. The true scope of the invention is intended to be embraced within the following claims.

We claim:

l. A process for treating a Viscous feedstock which is essentially a crude oil to effect both 1) hydroreiining and (2) reduction in viscosity to a predetermined pourpoint which is substantially lower than the pour-point of said feedstock, which comprises rst subjecting said feedstock to partial vaporization in equilibrium with a gas phase consisting essentially of hydrogen-rich recycle gas derived from the hydrorening step hereinafter defined, said partial vaporization being carried out at a temperature within abofut 100 F., and a pressure within about 300 p.s.i.g. of the temperature and pressure prevailing in the hydrorefining step hereinafter defined, thereby producing a gas phase comprising the .major proportion of said feed oil and a liquid phase comprising a minor proportion thereof, separating said liquid phase and subjecting the same to thermal coking to produce a coker distillate and coke, blending at least a portion of said Coker distillate with said gas phase, and then subjecting the resulting blend to hydrorefning in the presence of a hydroreiining catalyst at a temperature -between about 600 and 875 lF., a pressure between about 100 and 5000 p.s.i.g., a liquid hourly space velocity between about 0.5 and 15, and a hydrogen ratio between about 300 and 8,000 scf. per barrel of liquid feed, the conditions of said hydroretining being further correlated within the stated ranges so as to provide overall mild conditions incapable of effecting any substantial reduction in pour-point of the oil during hydrorening there-by providing an extended catalyst` life, cooling the product from said hydroreiining step and condensing the major portion thereof at substantially the pressure prevailing in said hydroreiining zone, thereby producing a supernatant hydrogen-rich recycle gas phase, contacting at least a portion of said recycle gas phase with said initial feed to effect partial vaporization as above described, recovering product oil of said predetermined pour-point from said condensing step, and further adjusting the temperature and hydrogen/oil ratio in said partial vaporization step, upwardly to increase vaporization and downwardly to decrease vaporization, so as to provide a liquid phase/vapor phase feed ratio therein which is inversely proportional to the desired pour-point of said product oil.

2. A process as defined in claim 1 wherein said portion of coker distillate comprises a fraction thereof boiling Ibelow about 700 F., and wherein the fraction of coker distillate boiling above about 700 F., is blended and recycled with the lfeedstock to said partial vaporization step.

3. A process as defined in claim l wherein said portion of coker distillate consists of the full-range depropanized bottoms thereof.

4. A process as defined in claim 1 wherein said hydrorening catalyst consists essentially of a minor proportion of cobalt oxide plus molybdenum oxide supported on a carrier which is predominantly activated alumina.

5. A process as defined in claim 1 wherein said feedstock is essentially crude -shale oil.

6. A process for treating a viscous feedstock which is essentially a crude oil to effect both (l) hydroreiining and (2) reduction in viscosity to a predetermined pourpoint which is substantially lower than the pour-point of said feedstock, which comprises first subjecting said feedstock to partial vaporization in equilibrium with a gas phase consisting essentially of (l) hydrogen-rich recycle gas derived from the hydrorening step hereinafter dened, and (2) fresh makeup hydrogen required for said hydrorening step, said partial vaporization being carried out at a temperature within about F., and a pressure within about 300 p.s.i.g. of the temperature and pressure prevailing in the hydroreiining step hereinafter defined, thereby producing a gas phase comprising the major proportion of said feed oil and a liquid phase comprising a minor proportion thereof, separating said liquid phase and subjecting the same to thermal coking to produce a coker distillate and coke, blending at least a portion of `said Coker distillate with said gas phase, and then subjecting the resulting blend to hydrorening in the presence of a hydroretining catalyst at a temperature between about 600 and 875 F., a pressure between about 100 and 5000 p.s.i.g., a liquid hourly space velocity between about 0.5 and 15, and a hydrogen ratio between about 300 and 8000 s.c.lf. per barrel of liquid feed, the conditions of said hydrorening being further correlated within the stated ranges so as to provide overall mild conditions incapable of effecting any substantial reduction in pour-point of the oil during hydrorefning thereby providing an extended catalyst life, cooling the product from said hydrorening step and condensing the major portion thereof at substantially theV pressure prevailing in said hydroreiining zone, thereby producing a supernatant hydrogen-rich recycle gas phase, contacting at leasty stock in said partial vaporization step is first equilibrated' with said recycle gas and the resulting liquid phase is then stripped with said fresh makeup hydrogen.

8. A process as detined'in claim 7 wherein the rate of injection of fresh makeup hydrogen in said stripping operation is substantially the same as the rate of hydrogen consumption in said hydroreiining step.

References Cited in the tile of this patent UNITED STATES PATENTS 2,606,141 MeyerV Aug. 5, 1952 2,691,623 Hartley Oct. 12, 1954 2,749,282 Porter et al. June 5, 1956 2,844,517 Inwood July 22, 1958 2,883,337 Hartley et al. Apr. 21, 1959 

1. A PROCESS FOR TREATING A VISCOUS FEEDSTOCK WHICH IS ESSENTIALLY A CRUDE OIL TO EFFECT BOTH (1) HYDROREFINING AND (2) REDUCTION IN VISCOSITY TO A PREDETERMINED POURPOINT WHICH IS SUBSTANTIALLY LOWER THAN THE POUR-POINT OF SAID FEEDSTOCK, WHICH COMPRISES FIRST SUBJECTING SAID FEEDSTOCK TO PARTIAL VAPORIZATION IN EQUILIBRIUM WITH A GAS PHASE CONSISTING ESSENTIALLY OF HYDROGEN-RICH RECYCLE GAS DERIVED FROM THE HYDROREFINING STEP HEREINAFTER DEFINED, SAID PARTIAL VAPORIZATION BEING CARRIED OUT AT A TEMPERATURE WITHIN ABOUT 100*F., AND A PRESSURE WITHIN ABOUT 300 P.S.I.G. OF THE TEMPERATURE AND PRESSURE PREVAILING IN THE HYDROREFINING STEP HEREINAFTER DEFINED, THEREBY PRODUCING A GAS PHASE COMPRISING THE MAJOR PROPORTION OF SAID FEED OIL AND A LIQUID PHASE COMPRISING A MINOR PROPORTION THEREOF, SEPARATING SAID LIQUID PHASE AND SUBJECTING THE SAME TO THERMAL COKING TO PRODUCE A COKER DISTILLATE AND COKE, BLENDING AT LEAST A PORTION OF SAID COKER DISTILLATE WITH SAID GAS PHASE, AND THEN SUBJECTING THE RESULTING BLEND TO HYDROREFINNG IN THE PRESENCE OF A HYDROREFINING CATALYST AT A TEMPERATURE BETWEEN ABOUT 600* AND 875*F., A PRESSURE BETWEEN ABOUT 100 AND 5000 P.S.I.G., A LIQUID HOURLY SPACE VELOCITY BETWEEN ABOUT 0.5 AND 15, AND A HYDROGEN RATIO BETWEEN ABOUT 300 AND 8,000 S.C.F. PER BARREL OF LIQUID FEED, THE CONDITIONS OF SAID HYFROREFINING BEING FURTHER CORRELATED WITHIN THE STATED RANGES SO AS TO PROVIDE OVERALL MILD CONDITIONS INCAPABLE OF EFFECTING ANY SUBSTANTIAL REDUCTION IN POUR-POINT OF THE OIL DURING HYDROREFINING THEREBY PROVIDING AN EXTENDED CATALYST LIFE, COOLING THE PRODUCT FROM SAID HYDROREFINING STEP AND CONSENSING THE MAJOR PORTION THEREOF AT SUBSTANTIALLY THE PRESSURE PREVAILING IN SAID HYDROREFINING ZONE, THEREBY PRODUCING A SUPERNATANT HYDROGEN-RICH RECYCLE GAS PHASE, CONTACTING AT LEAST A PORTION OF SAID RECYCLE GAS PHASE WITH SAID INITIAL FEED TO EFFECT PARTIAL VAPORIZATION AS ABOVE DESCRIBED, RECOVERING PRODUCT OIL OF SAID PERDETERMINED POUR-POINT FROM SAID CONDENSING STEP, AND FURTHER ADJUSTING THE TEMPERATURE AND HYDROGEN/OIL RATIO IN SAID PARTIAL VAPORIZATION STEP, UPWARDLY TO INCREASE VAPORIZATION AND DOWNWARDLY TO DECREASE VAPORIZATION, SO AS TO PROVIDE A LIQUID PHASE/VAPOR PHASE FEED RATIO THEREIN WHICH IS INVERSELY PROPORTIONAL TO THE DESIRED POUR-POINT OF SAID PRODUCT OIL. 